Hydrocarbon conversion process



. Feb. zs, 1965 L. J. SPILLANE ETAL HYDROCARBON CONVERSION PROCESS Filed Sept. 50, 1960 FIGURE l 2 Sheets-Sheet l III IGNITER 5 REACTION CHAMBER EXHAUST CHAMBER TAKE-OFF FOR LIQUID PRODUCT 7,,ALTERNATE 0 OR AIR ENTRY TAKE OFF FOR NON-CONDENSED MATERIALS .i COOLANT ENTRY IN VEN TOR-S.

LEO J. SPILLANE BY ROY L. GRANTOM ATTORNEY 1965 L. J. SPILLANE ETAL 3,170,863

HYDROCARBON CONVERSION PROCESS Filed Sept. 30, 1960 2 Sheets-Shet 2 FIGURE 2 $83222 ggg 23 --PREHEATER 23, ,PREHEATER 3A'--2 L i sxnms CHAMBER CONDENSERS INVENTOR. LEO J; SPILLANE BY ROY L. GRANTOM United States Patent O assignors to Monsanto Company, a corporation of' Delaware Filed Sept, 30, 1960, Ser, No. 59,620 9 Claims. (Cl. 208-3) The present invention relates to a process for the thermal transformation of hydrocarbon mixtures. More particularly, the present invention relates to a process for the thermal transformation by non-catalytic partial oxidation of a liquid hydrocarbonmixture into a liquid fraction of an increased aromaticity and a gaseous fraction comprising low-molecular weight unsaturated hydrocarbons.

It is an object of this invention to provide a process for the thermal transformation of liquid hydrocarbon mixtures. A further object of this invention is to provide a process whereby liquid hydrocarbon mixtures are converted by non-catalytic partial oxidation into a liquid fraction of increased aromaticity and a gaseous fraction comprising low-molecular weight olefinic hydrocarbons. Another object of this invention is to provide a process for the non-catalytic partial oxidation of hydrocarbon mixtures boiling above 200 C. whereby gaseous olefin hydrocarbons and liquid fractions within the boiling range of gasoline are produced at significantly lower temperatures than other thermal methods. A still further object of this invention is to provide a process for the production of gasoline fractions of increased aromaticity and octane value. Another object of this invention is to provide a process for the non-catalytic partial oxidation of liquid hydrocarbon mixtures into gaseous olefins wherein there is produced a high ratio of the higher molecular weight gaseous olefins to the lower molecular weight gaseous olefins. Yet another object of this invention is to provide a process for the non-catalytic partial oxidation of liquid hydrocarbon mixtures which will operate at significantly lower temperatures than previous processes. Additional objects will become apparent from the description of the invention herein disclosed.

In fulfillment of these objects, it has been found that r when liquid hydrocarbon mixtures with boiling ranges of 65 to 400 C. and higher are subjected to partial com bustion in a properly designed single stage reaction chairiher having a constricted opening at the exit end at tem-. peratures of 375 to 700 C. and pressures of 25 to 100 p.s.i.a. and a ratio of oxygen to hydrocarbon mixture of 0.01 to 2.0 mols of oxygen per mol of hydrocarbon mixture and at relatively low flow rates and residence time in the reaction chamber, a liquid fraction of in creased aromaticity and octane number is obtained together with a gaseous fraction comprised of gaseous olefins of 2 to 4 carbon atoms. Small amounts of oxygen ated products such as acrolein are also obtained in the product.

The apparatus and operation of this invention is more aptly described and illustrated by reference to the accompanying drawings. FIGURE 1 presents a cross section of the reaction unit which is preferred in the practice of this invention. FIGURE 2 is a flow diagram of the entire process including the reaction unit of FIGURE 1.

Referring to FIGURE 1, a preheated liquid hydrocarbon feed stream is passed into the mixing head 2 by means of line 1. Preheated oxygen or air is brought into the mixing zone throughline 3. The mixed hydrocarbons and oxygen are then forced into the reaction chamber 5 through a spray nozzle 4 which increases uniformity of mixing of the feed fuels entering into the reaction chamber. tage of preventing flashbacks, premature ignition, etc. The hydrocarbon and oxygen sprayed into the reaction Thespray nozzle has the added advan-' Patented Feb. 23, 1965 chamber are ignited by an igniting surface 6. The igniter may be a spark plug or any other suitable device and may be placed anywhere in the reaction chamber but must remain so located that the incoming oxygen and hydrocarbon will strike the igniting surface. Once combustion is initiated the igniting surface is of no further use since the reaction'is autogenous. It is not neces sary that an igniting surface be used since spontaneous ignition will occur if the oxygen and feedstock are heated prior to entry into the reactor.

In FIGURE 1 is shown an alternate arrangement for the introduction of the oxygen or air into the reaction chamber. This alternate arrangement allows the heat of the reaction chamber to be used for preheating purposes. This is accomplished by introducingthe oxygen into the reaction chamber through line 7 and annular passage 8. If this alternate arrangement is not used, the annular passages are closed off at the elbow 9 and'thus they be come insulating aids.

The reaction chamber may be lined with any suitable refractory material. It also maybe of any shape and size consistent with good flow characteristics and certain limitations as to length and Width discussed later in the specification. The reaction chamber of this process is characterized by having a constricted opening at its exit end, the optimum amount of constriction varying with feedstocks and other considerations.

The reaction product of the combustion of the hydrocarbon feed in the reaction chamber next passes through the constricted opening 11. Generally, for flow and other design characteristics, it is desirable to use a converging and diverging nozzle arrangement as exemplified by the sloping lines leading to the constricted opening. It is to be understood, of course, that the degree of slope and the ratios of slope of convergence to slope of divergence are not to be in any way limited by the drawing in FIG- URE l. On passing through the constricted opening the gases are expanded into the exhaust chamber 12. The rapid expansion of the gases in the exhaust chamber brings about a proportionally rapid lowering of the temperature of the gaseous reaction products. The diver gence of the exhaust chamber may be so designed that the expansion of the hot product gases Will sufficiently quench the effluent gaseous reaction product stream thus alleviating any need for further cooling. However, a coolant such as water may be injected into the exhaust stream to aid the expansion quench. A coolant has the added advantage of slowing the high velocity gaseous products and simultaneously removing the heat generated by the slowing action. This coolant may be injected through line 13 or may be injected through the walls at an angle tothe path of the exhaust stream or may be injected co-current with the exhaust stream. The liquid products which are condensed from the exhaust stream are collected in tank 15 and are removed by line 16. The lighter non-condensed gases exit the exhaust chamber through line 14-.

FIGURE 2 presents the flow arrangement of the process of this invention. The hydrocarbon mixture storage ,or source is represented by 21. Feed from storage flows by line'22 into a preheater 23 and then by line 24 into the mixing zone 25. The oxygen or air is taken from storage or source 26 by means of line 27 into preheater 28 and then by line 29 into the mixing zone 25. If a diluent gas such as steam, nitrogen, etc., is used, it may be introduced into the oxygen line through line 30. Also, if it is desired to use the alternate arrangement of oxygen entry described in the explanation of FIGURE 1,

V the oxygen stream may be fed into the reaction chamber the exhaust chamber 33. In this chamber cooling of the hot products takes place, either by expansion of the hot gases or by both expansion and injection of a coolant or by injection of a coolant alone. The liquid products collect in tank 34 and are removed by line 35. The liquid products then may be removed through line 37 or line 36. Non-condensed product material flows from the exhaust chamber by line 38 into a second condensing unit where additional liquid product is condensed from the product stream. From this second condenser, the gaseous product flows to a third condenser 42 by means of line 41 and the still non-condensed gases exit into still a fourth condensing unit 45 through line 44. From this final condenser, a non-condensed gaseous product comprising low molecular weight olefins flows by line 47 to its future utility. Lines 40, 43, 46, represent the liquid product removal lines from the secondary condensing system. The number of condensers shown in the drawing is not limiting since this is entirely dependent upon the efliciency of the expansion quench and the secondary condenser or condensers and the overall degree of condensation desired. The drawing indicates that all liquid product may eventually be gathered in a single line. However, this may not always be desired since by regulating the temperatures of the three condensing units a certain degree of fractional distillation may be obtained and it may be preferred to keep the fractions from each unit separate. If it is preferred to pass the liquid product fraction through a distillation unit or to recycle a portion of this fraction, line 37 would be used to move the liquid to such future processing.

It will become readily apparent to those skilled in the art that a great number of variations and modifications of the equipment and flow arrangement as presented in FIGURES 1 and 2 may be made without departing in any way from the spirit and scope of this invention. It is to be understood that within the preceding process description the use of the word oxygen is to include air also.

To further illustrate the invention herein disclosed, the following examples are presented. It is to be understood, of course, that these examples are in no way to be construed as limiting the application, operation, or conditions of this invention.

Example I The feedstock to this run was a 93 to 204 C. naphtha containing approximately 12 weight percent of aromatics. Its research octane number (RON) was 26 clear and 57 with 3 cc. of tetraethyllead. The reaction unit was constructed similar to FIGURE 1 of stainless steel coated on the inside with Norton Alundum. The reaction chamber length was 4 inches and the diameter 1% inch. The opening at the exit end was 0.1 inch in diameter and the angle of the converging slope 60 and that of the diverging slope The oxygen source was air. Operating conditions for the run were as follows:

Naphtha flow rates, cc./min. 175 Air flow rates, liters/min. 55 Reactor chamber temperature, C. 525-527 Reactor chamber pressure, p.s.i.g 59 Chamber flow velocity, ft./sec. 3.17 Chamber residence time, sec. 105.6 10

The yield and product data were as follows:

Liquid recovery, weight percent of feed 85.5 Aromatics, weight percent of liquid product 15.22 Clear Research Octane No. 54-58 It is readily apparent that the liquid product contains approximately a percent increase in aromatic concentration over that of the feed material. Further, the unleaded octane number of the liquid product represents a 115 percent increase over that of the feed material.

4 Example 11 The feedstock and reactor used in this test were the same as those of Example I. In this run steam was 1njected with the air. Operating conditions were as follows:

'Naphtha flow rate, cc./min. 108 Air flow rate, liters/min. 24 Water flow rate, cc./min. 41 Naphtha entry temperature, C 540-595 Air-steam entry temp., C. 425-515 Reactor chamber temperature, C. 580-610 Reactor chamber pressure, p.s.i;g. 52-55 Chamber flow velocity, ft./sec. 4.25 Chamber residence time, sec. 48.6 10

The water was used as a diluent to the air stream, and though metered as liquid was used as steam.

The yield and product data were as follows:

Liquid recovery, weight percent of feed 69.5 Distribution, volume percent of liquid product:

Parafiin-naphthene 71.91 Aromatics 23.09 Oxygenated compounds 5.00 Gaseous olefins recovered, weight percent of feed:

Ethylene 4.9 Propylene 3.6 Butylene 4.3 Methane recovery, weight percent of feed 2.9 Research Octane No. liquids (clear) 51.2 Research Octane No. (3 cc. TEL) 73.8

The ocane number for the liquid portion of the product from this partial oxidation of the 93-204 C. naphtha represent a 97.0 percent increase in the clear octane No. and a 46.5 percent increase in the leaded value. The aromatic content of the liquid product represents a 92.6 weight percent increase in aromatics over that of the feed material. Also, it should be noted that the amount of butylenes produced is nearly equal to the amount of ethylene produced.

Example III This example demonstrates the present invention as applied to a catalytic cycle oil feedstock. Two runs were made, each under difierent operating conditions. In run No. 1 more severe operating conditions were utilized producing a larger amount of gaseous than liquid product. Run No. 2, however, shows the effect of milder operating conditions whereby the primary aim is to produce liquid products.

The feedstock to this example was a 210-330 C. catalytic cycle oil containing approximately 55 weight percent aromatic and unsaturated compounds, the aromatic compounds being both mono-, and polycyclic and being poly substituted with relatively short chain alkyl The reactor used for the following runs was constructed as illustrated in FIGURE 1 with the cylindrical reaction chamber being one inch in diameter and 5% inches in length. The reactor was lined with Norton Alundum. The opening at the exit end was 0.14 inch in diameter with a converging slope of 30 and a diverging slope of. 7.5. The oxygen source was air.

ans-asst; V

Run No. 1 Run No. 2

Oil flow rate, celmin 91 41 Air flow rate, liters/min 99 34 Oil entry temperature, C 362-384 391-433 Air entry temperature, C 541-571 544-555 Reactor temperature, O. 527-638 462-487 Reactor chamber pressure, p. 30 18-19 Chamber flow velocity, it./sec 7. 4 3. 18 Chamber residence time, sec 61. 8X10- 142. 8X10- The yield and product data were as follows:

Liquid recovery, weight percent of feed--. 56. 7 97.7 Weight'percent liquid in gasoline range 7. O 13. 5 Gaseous; olefin recovery, weight percent of d .16. 9 7. 0

In the practice of the invention herein disclosed, feedstocks ranging from light naphthas to relatively heavy crudes may be partially oxidized to yield a liquid fraction of a higher degree of aromaticity than the feedstock and also gaseous olefinic hydrocarbons. The light naphtha fractions are generally those whose average boiling point may be as low as 65 C. The other extreme, a relatively heavy crude may have a boiling range of 275 to 400 C. and higher. depend somewhat on the product desired in that the lower the initial boiling point of the charge stock the lower the liquid recovery obtained. The feedstocks useful in the practice of this invention must contain some aromatic material in order to obtain the product of greater aromaticity. Little or no aromatics are actually formed in the process of this invention. The present invention causes the selective cracking away of the alkyl substituents to the aromatic nucleus and also cracking of straight, branched and cyclic parafiins. Whenoperated at the preferred temperatures very little or no rupturing of the aromatic nucleus takes 'place. The amount of aromatics desired in the feedstocks may range as low as 1 percent. However, the desired aromatic content will depend .on whether aromatic concentration is the primary objective or whether the production of olefinic gases is primary. Obviously, if products of high aromaticity and octane value are preferred with gaseous olefin production secondary, then feedstocks of relatively high aromatic content would be preferred. If low aromatic feedstocks are used, the product may be recycled over and over until the desired aromaticity is obtained; However, it is not necessary that all of the octane improvement be obtained from the concentration" of aromatics since olefins in the liquid product will also increase the octane value.

The term aromaticity is used throughout this specification to denote relation to the aromatic nucleus. More specifically, it is used in the-sense that the closer the compound approaches benzene or naphthalene or other aromatic nucleus, the greater is the aromaticity. For example, toluene would be of greater aromaticity than propylbenzene, xylene'of greater aromaticity than dipropylbenzene, etc. I r

The reaction temperatures which are suitable for the practice of the invention herein disclosed may range from 375 to 700 C. A preferred temperature range for the reaction is from approximately 450 to 600 C., however,

the optimum temperature range will vary according to the feedstock and with the degree of cracking desired. Generally, reaction temperatures of 450 to 500 C. produce the greatest yield of liquid product in the gasoline boiling range and temperatures around 600 C. produce the greatest yield of ethylene, propylene, etc., without also producing considerable carbon formation. At temperatures in excess of 625 C. the'yield of ethylene increases but there is substantial carbon formation and :at temperatures in excess of 700 C. there is almost complete destruction of the charge to carbon, carbon monoxide, carbon dioxide and water. At the 400 to 500 C. reaction temperature, the alkylside chains split off from the aromatic portion and littlefurther cracking of the detached side chain; As the temperature is increased the detached The boiling point of the feedstock will a 6 side chains and the paraflin and naphthene hydrocarbons become increasingly susceptable to further cracking, resulting in increased ethylene yield. A further increase in temperature to above 625 C. may result in the rupturing of the aromatic nucleus resulting in greater carbon formation. If the temperature in the reaction chamber is allowed to fall below approximately 375 C. which is the upper limit of the 300 to 375 C. cool flame region, considerable quantities of oxygenated products will be produced. The temperatures at which this process is operated are a significant advantage of the present invention in that they are substantially lower than those generally known to the art for thermal conversion processes producing similar products.

Temperatures in the reaction chamber may be controlled by controlling preheat temperature, feed rates, feed to oxygen ratios, and residence time in the reaction chamber. 'Diluent gases also aid in reactor temperature control. In general, the temperature of the reactor lags behind the preheated streams until a reactor temperature of 325 to 375 C. is reached. At this point, if no ignitor is used spontaneous ignition occurs. The temperature of the reactor surges rapidly, almost instantaneously, to a new temperature of 450 to 475 C. All other variables remaining fixed the rate of increase of reactor temperature then decreases somewhat and proceeds to follow roughly the range of increase of preheat temperature. The

eifect of feed rates and ratios and residence time in the reactor chamber will be discussed at some length later in the specification.

The effect of reaction chamber pressure is somewhat analogous to that of temperature. This is to say that in a given temperature range an increase in chamber pressure will result in an increase in the yield of the lower molecular weight, gaseous olefins. This is the reverse of the result expected based on the thermodynamic considerations involved. When lower boiling feedstocks such as the 93204 C. naphtha of Example I are used, the yield of gaseous olefinic hydrocarbons is still increasing at p.s.i.g. However, for a heavier feedstock such as the 210-330" C. burner oil of Example 111, there is an increase in gaseous olefins produced until approximately 6 0 p.s.i.g. is reached then there is a rather rapid drop olf in production of these gaseous products. These facts would indicate once more that the optimum operating conditions will depend largely on the feedstock and on whether production of the liquid or the gaseous product is the primary objective. However, it has been found that adequate results'with any of the feedstocks within the scope of this invention may be obtained within a range of chamber pressures of from 25 to p.s.i.a., witha more preferred range being from 40 to 65 p.s.i.a. It is to be understood, of course, that these ranges, though generally. encompassing the most practical and desired pressure ranges are not to be taken as limiting the present invention in all of its scope.

Another variable which has considerable effect on the efliciency of the process herein disclosed is the hydrocarbon feedstock-oxygen ratio. At any given temperature and chamber pressure the yield of various gaseous hydrocarbons is greatly dependent upon the oxygen concentration in the reactor. The yields of gaseous, olefinic hy- :drocarbons increases as the concentration of oxygen is increased until a maximum of approximately 50 mol percent oxygen is obtained in the reaction chamber. At this point there begins a decrease in the yields of gaseous .olefins. Generally, throughout this application reference is made to oxygen, however, it is to be understood that this does not limit'the practice of the present invention to pure oxygen. On the contrary, it is more practical and much less expensive to use air. Of course, when air is used, it is necessary to take into consideration the inert components of the air in determining flow rates, concentrations, heat requirements, etc. Though it has been determined that some cracking will take place in the present process in the complete absence of oxygen, this does not occur to any extent. At 10 mol percent oxygen concentration the yields of low molecular Weight unsaturated hydrocarbons is still not appreciable but as the oxygen concentration is increased beyond this point, yields of olefinic gases increase rather rapidly. In its broad application or in some specific situations the oxygen concentration may range anywhere within the broad limits of mole percent to 80 mol percent. However, a more practical and preferable range of oxygen concentration is to 60 mol percent with a still more preferred range being from to 50-mol percent oxygen. It has been found advantageous in many situations to inject steam into the oxygen containing stream. Steam injection aids in maintaining oxygen concentrations without dilution of the gaseous products with a non-condensible material.

Input rates of both the hydrocarbon mixture and the combustion supporting gases have considerable eifect on the reactor temperature. As input flows increase there is an increase in temperature. This, of course, may be compensated for to some degree by the manipulation of the other operating variables. The feed flow rates which are within the practical applicability of the present invention, as set forth herein, are 0.5 to 10 volumes of feed per minute per volume of reaction space within the reaction chamber. More preferred flow rates for the hydrocarbon feed are found in the range 1.0 to 4.0 volumes of feed per minute per volume of reaction space. The input flow rates of the combustion supporting gases and diluents will, of course, be dependent upon the desired ratio of these gases to the feed mixture. The effect of the feed input flow rate on the product is as would be expected. If high feed flow rates are used with the result of higher temperatures, this will cause the formation of more gaseous olefins. However, if the temperature is held constant by means not affecting the operation other than as temperature controls, e.g., diluent gases, then increased feed flow rates will result in a greater liquid yield of improved aromaticity.

The product change due to feed flow rate changes, other than being a result of the corresponding temperature change, is caused by the change in residence time and flow velocity within the reaction chamber. Short residence time at given conditions will produce more liquid product and less gaseous product. Conversely, long residence time under the same given conditions will cause a greater gaseous olefin product. Residence time under any particular feed flow rate will be limited by temperature, pressure, feedstock, reactor flow velocity, and the reactor design. The last of these, reactor design is fixed and is not, therefore, considered a variable effecting residence time. Residence time may generally vary from l0 10 seconds to 250x10 seconds. A somewhat more preferred range of residence time is 50 l0- sec onds to 200 l0 seconds. Optimum residence times as previously stated, will be dependent on temperatures, pressures, feed flow rates, reactor flow velocities, feedstocks, desired product, and reactor design.

The flow velocity of the reactants through the reaction chamber should be within the broad range of from 1 to 15 feet per second. Velocities of 2 to 10 feet per second will generally yield somewhat better results, however. Within these ranges flow velocities will vary in accordance with the considerations expressed in regard to residence time, e.g., feed flow rates, temperatures, pressures, feedstocks, desired product and reactor design.

The reaction unit utilized in the process herein described will generally have a cylindrical internal surface. It may be constructed of any material which will withstand the physical requirements of temperature and pressure and will not adversely effect the reaction taking place within. Materials adversely effecting the reaction may be used with an internal lining of a suitable refractory material such as Alundurn, hard carbon or graphite. A

8 non-limiting example of a suitable reaction unit is one constructed of steel with an internal fusedalumina lining.

In the design of the reaction chamber it is of some importance that a proper ratio between the length and diameter of the reaction chamber be maintained. This is especially true in processing heavier feedstocks in that if the reaction chamber is overly long in relation to its diameter, a significant amount of carbon and tar formation takes place as a result of secondary condensations in the reaction chamber. By utilizing a proper reactor length much or all of this carbon and tar formation caused by secondary condensations may be eliminated. The criticality of the length to diameter ratio becomes less as the feedstock becomes lighter or lower boiling. The ratio of length to diameter may vary from 0.5:1 to 10:1, however, it is somewhat more desirable to use ratios of 1:1 to 4:1.

The relationship which must exist between the diameters of the constricted exit orifice of the reaction chamber and the reaction chamber is a critical factor in the design of the reaction unit. This relationship has considerable effect on reactor flow velocities, residence time and pressures in that the pressure ratio between the reaction chamber and the exhaust or expansion chamber is primarily controlled by the size of the exit orifice. For this reason the ratio of the diameters of the exit orifice and the reaction chamber is generally expressed as a pressure ratio of the reaction chamber side of the orifice to the exhaust or expansion chamber side. The pressure ratios at which the present process is operable may range from 1.2:1 to 10:1. A more preferred pressure ratio, however, would be within the range of 2:1 to 7:1.

The exit orifice of the reaction chamber generally consits of the orifice itself and a converging and diverging section as shown in FIGURE 1. The angles of convergence and divergence are given as the angle between the converging and diverging slope and the plane of the walls of the reaction chamber. The angle of convergence from the reaction chamber wall to the exit orifice may range from 10 to with a generally more useful and practical range being from 30 to 75. The angle of divergence from the exit orifice wall to the exhaust chamber wall may vary from 5 to 90, a more preferred range being 30 to 75.

The exhaust or expansion chamber of this reaction unit is not strictly defined. It generally is considered to include' the diverging area immediately following the exit orifice. Actually, it is within this diverging section of the exit orifice that expansion begins and in many cases all expansion takes place in this area. The exhaust or expansion chamber then functions as an area where the gaseous products exiting the reaction chamber are rapidly expanded and thereby cooled. The gaseous products on passing through the exit orifice attain high velocities and in expansion in the diverging area of the expansion chamber reach even higher velocities. It is a second function of the exhaust or expansion chamber to decelerate these gaseous products back to speeds practical for condensation and collection of the vaporized liquid products and for control of the product gases.

The size of the exhaust chamber is not critical within broad limitations. It must be large enough in diameter to allow some expansion and of a length suificient for deceleration of the high velocity gaseous products. Water or other coolants may be injected into the hot high velocity gases to aid both in cooling and in slowing these gaseous products. Injection of coolants may be made counter-current co-current, or at an angle to the direction of fiow of the gaseous products. The use of coolants .will generally allow smaller exhaust chambers in the reaction units. The upper limit in size of the exhaust chamber is governed only by the practical physical size of the chamber.

The exhaust chamber as shown in the drawing, FIG- URE 1, provides a collection tank for condensed liquids at its end. While this is a very useful arrangement for collection of the product condensed within the exhaust chamber it is not intended to be limiting on the present invention in any way. Any number of methods for collecting the liquid products condensed in the exhaust chamber will be readily apparent to anyone skilled in the art.

For introduction of the reactants into the reaction chamber, a spray type mixing head as illustrated by FIGURE 1 is preferred. However, any of the many variations of this type of mixing head or any other mixing head may be used in the practice of the present invention, though not necessarily with equivalent results. The reactants may also be. introduced separately as by the introduction of one of the reactants through annular passages surrounding the reaction chamber as shown in FIGURE 1 of the accompanying drawings. This method of introducing reactants into the reaction chamber has the advantage of utilizing heat produced in the reaction chamber for the preheating of the reactant.

In the practice of the present invention it is often found to be advantageous to use diluents in the reaction chamber. The diluentsmay be water (generally as steam), nitrogen, or other such inert materials. Diluents function to reduce carbon and coke formation by preventing polymerization and other secondary reactions of the olefins formed in the reaction chamber. Also, diluents aid in the control of the temperature Within the reaction chamber. Probably the most useful diluent, from a process standpoint, is steam because of its higher heat capacity and the ease with which it can be separated from the products.

What is claimed is:

1. A process for the non-catalytic partial oxidation of liquid hydrocarbon feedstocks containing at least one percent of aromatic hydrocarbons to produce a liquid fraction of increased aromaticity and a gaseous fraction comprised of olefinic hydrocarbons of 2 to 4 carbon atoms, the process'comprising introducing the liquid hydrocarbon feedstock concurrently with an oxygen containing gas selected from the group consisting of oxygen, air and mixtures thereof in a mol ratio of oxygen to hydrocarbon feedstock of 0.1:1 to 2:1 into a reaction zone at a hydrocarbon feed input rate of 0.5 to 10.0 liquid volumes of feed per minute per volume of reaction space within the reaction zone, the temperature within the reaction zone being in the range of about 375 to 700 .C., the reaction zone pressure being in the range of 25 to 100 p.s.i.a., the flow velocity of the reactants within the reaction zone being approximately 1 to 15 feet per second and the residence time of the reactants in the reaction chamber being in the range of 10 10 seconds to 250 10- seconds, thereafter passing the high temperature gaseous products into a cooling zone wherein said gaseous products are rapidly cooled by adiabatic expan sion' of the gaseous products into the cooling zone, the cooling zone being maintained at a pressure with respect to the pressure of the reaction zone such as to cause a reaction zone to cooling zone pressure ratio of 1.2:1 to 10:1.

2. The process of claim 1 wherein the liquid hydrocarbon feedstock is a naphtha having a boiling range of 90 to 210 C.

3. The process of claim 1 wherein the mol ratio of oxygen to hydrocarbon feed is in the range of 02:10 to 10:10.

4. The process of claim 1 wherein the pressure in the reaction chamber is 40 to 65 p.s.i.a.

5. The process of claim lwherein the pressure ratio I between the reaction zone and the cooling zone is mainoxygen contain- Refereuces Cited in the file of this patent UNITED STATES PATENTS 2,431,515 Shepardson Nov. 25, 1947 2,661,380 Orkin DecQl, 1953 2,692,292 Robinson Oct. 19, 1954 2,722,553 Mullen et al. Nov. 1, 1955 2,750,434 Krejci June 12, 1956 2,775,629 Anderson Dec. 25, 1956 2,823,243 Robinson Feb. 11, 1958 2,945,074 Elliott et al. .a July 12, 1960 2,970,178 Braconier et a1 Jan. 31, 1961 3,018,309 Krejci Jan. 23, 1962 3,049,574 Johnson Aug. 14, 1962 

1. A PROCESS FOR THE NON-CATALYTIC PATRIAL OXIDATION OF LIQUID HYDROCARBON FEEDSTOCKS CONTAINING AT LEAST ONE PERCENT OF AROMATIC HYDROCARBONS TO PRODUCE A LIQUID FRACTION OF INCREASED AROMATICITY AND A GASEOUS FRACTION COMPRISED OF OLEFINIC HYDROCARBONS OF 2 TO 4 CARBON ATOMS, THE PROCESS COMPRISING INTODUCING THE LIQUID HYDROCARBON FEEDSTOCK CONCURRENTLY WITH AN OXYGEN CONTAINING GAS SELECTED FROM THE GROUP CONSISTING OF OXYGEN, AIR AND MIXTURES THEREOF IN A MOL RATIO OF OXYGEN TO HYDROCARBON FEEDSTOCK OF 0.1:1 TO 2:1 INTO A REACTION ZONE AT A HYDROCARBON FEED INPUT RATE OF 0.5 TO 10.0 LIQUID VOLUMES OF FEED PER MINUTE PER VOLUME OF REACTION SPACE WITHIN THE REACTON ZONE, THE TEMPERATURE WITHIN THE REACTION ZONE BEING IN THE RANGE OF ABOUT 375 TO 700*C., THE REACTION ZONE PRESSURE BEING IN THE RANGE OF 25 TO 100 P.S.I.A., THE FLOW VELOCITY OF THE REACTANTS WITHIN THE REACTION ZONE BEING APPROXIMATELY 1 TO 15 FEET PER SECOND 